Recovery of hydrogen from refinery gas



Aprll 19, 1966 A. RAMr-:LLA 3,246,951

RECOVERY OF HYDROGEN FROM REFINERY GAS Filed May 28, 1962 2 Sheets-Sheet1 Amilcore Rcmello ATTO/BVD A. RAMELLA April 19, 1966 RECOVERY OFHYDROGEN FROM REFNERY GAS 2 Sheets-Sheet 2 Filed May 28, 1962 INVENTORAmilcare Ramello ATM/wf] United States Patent 3,246,951 i RECOVERY FHYDRQGEN FROM REFINERY .GAS Amilcare Ramella, Woodbury, NJ., Yassignerto Socony Mobil Gil Company, Inc., a corporation of New York Filed May28, 1962, Ser. No. 197,987 14 Claims. (Cl. 23-212) The present inventionrelates rto .the recovery `of hydrovgen of high purity, of at least 90%`by volume, from industrial gas streams comprising hydrogen andcomponents not reactive with extracting .agent and, more particularly,to the recovery of hydrogen from refinery ygas streams containing atleast 20 percent by volume .of hydrogen and comprising hydrogen andhydrocarbons with or without ammonia, hydrogen .sulfide and the like.

Reforming of hydrocarbons and particularly petroleum hydrocarbonsinvolves dehydrogenation of naphthene, dehydrogenation of paraffins,dehydrocyclization of paraffins, isomerization of normal parains toiso-parans, and hydrocracking of paraliins. of reforming conditions the`greater the dilution -of hydrogen produced -in the dehydrogenation ofnaphthenes and the dehydrocycliza-tion of parat-:fins with hydrocarbonfragments such as methane, ethane, and propane.

With the increased demand for high octane gasoline, e.g., C+ reformatehaving an octane rating l(Research +3 -ml. TEL) Vof 100 the purity of:reformer hydrogen has decreased. 4Present demands for hydrogen on theother hand have increased.

Thus, for example, it is a matter yof economics whether MidContinentnaphtha be hydrogenated :to reduce the sulfur content to providereformer feed for a reformer unit employing particle-form platinum-groupmetal reforming catalyst. On the other hand, when thermal naphtha is tobe reformed in the presence of platinum-group :metal reforming catalystthere is no choice. The nitrogen content of thermal naphtha is so highthat the platinum-group metal reforming catalyst is inactivated in animpractically short time when the thermal naphtha is nothydrodesulfurized and hydrodenitrogenized prior to reforming.Consequently, the increased demand for upgrading previously acceptedgasolines has krequired reforming of `thermal naphtha and consequentlyincreased demand for hydroge to hydrodecontaminate the thermal naplitha.

It is now rather general practice to hydrodecontaminate domestic `fueloil and similar hydrocarbon fractions `boiling above gasoline to upgrade`these fractions thus making a further demand for hydrogen.

The problem of meeting the more stringent specifications for jet fuelsis being solved at least in part by hydrogenation of the naphthenes ofthe jet fuel.

The treatment of lubricating voil fractions in the presence of hydrogenhas also been added to the list of refinery operations using hydrogenproduced in the reforming unit. ln fact, the demand for hydrogen forpretreating reformer feed, hydrodesulfuring domestic fuel oil,hydrogenating jet fuel fractions, hydrotreating lubricatingfractions,and hydrocracking crudes, deasphalted crudes, topped crudes, and gas oilfractions exceeds the supply with the consequent result that thereformer make-gas is passed from one refinery operation to anotherprovided the contaminants picked up in one operation are not deleteriousto the succeeding operation.

Hydrogen can be produced by partial oxidation of hydrocarbons and byother means but not at .a cost which makes its use attractive in theaforementioned and other retinery hydroprocesses. On the `the hand,there are refinery gas streams containing hydrogen but `of such lowconcentrations that the use thereof is not practical.

For example, the off-gas of catalytic cracking of gas The greater theseverity 3,246,951 Patented Apr. 19, 1966 oil such 'as the olf-,gas fromthe Thermofor Catalytic Cracking of gas oil 4has the followingcomposition:

In other words, were'ICC olf-gas to 'be used in hydro.- processing 'forevery 1000 s.c.f. of hydrogen circulated 2855 s.-c.f. ,of diluent gasmust also be compressed and transported. Many of the hydroprocesses canbe carried out at lower pressure. That is .to say, .-it is the partialpressure of hydrogen not the -total pressure which is controlling inmany hydroprocesses Hence, the higher the purity of vthel'hydrogen-containing gas the lower the .total pressure for a givenhydrogen partial pressure. The lower r the pressure the lower thecapital cost; the higher the purity .of the hydrogen-containing gasstream .the lower the compression costs; the higher .the purity of thehydrogen-containing gas the lower the aging rate andthe longer theinterval between regenerations, i.e., the longer the onstream .time andthe fewer the off-stream Ytimes for regeneration .of the catalyst in agiven calendar interval. For example, a decrease .of about 20% ininvestment cost of a new hydrodesulfurization unit treating middlekdistillate .can `be realized by employing recovered hydrogen of .90%purity over `the Vuse of .reformer make-gas of `60% purity. For ahydrodesul'furization unit treating 7,000 barrels .of middle distillateper day the saving in investment cost .can amount to about .$250,000. Onan .existing hydrodesulfurization unit employing lthe recovered hydrogenof purity can double the on-stream time of the unit betweenregenerations.

Th-us, it is manifest that a means of raising the hydrogen `content ofrefinery gas streams will make available hydrogen which cannot bepresently used and the availability of 90 Vpercentnby volume hydrogenwill present additional advantages. The present invention provides thesolution tothe twin problems.

In general, the ,present invention provides for contacting any refinery.stream containing at least 20 percent .by volume of hydrogen with ahydrogenatable extracting agent in the presence of a hydrogenationcatalyst under hydrogenating conditions of temperature, pressure, andliquid hourly space velocity to react the hydrogen in the gas stream.with the hydr-ogenatable extracting agent to prolduce fat extractingagent and off-gas of reduced concenration of hydrogen. The fatextracting agent is separated from thedenuded gas stream. The fatextracting 4agent is then contacted with dehydrogenating catalyst underdehydrogenating conditions of temperature, pressure, and liquid hourlyspace velocity to produce a gaseous fraction containing at least 90percent by volume of hydrogen and :lean extracting agent. The highpurity gas stream flows lto .other hydroprocesses while the leanextracting 'agent is returned to the hydrogenation stage.

"The method of recovering hydrogen of at least 90 percent purity.briefly and generally described hereinbefore is to be distinguishedover the method ofpurifying reformer `recycle Agas described in U.S.Patent No. 2,328,828 issued in 1943.

The patentee describes a process in which naphtha having `an ASTM knockrating of about 25 `to 45 mixed with about 0.5 .to 4 molsof hydrogen permol of said naphtha is contacted with dehydrogenating catalyst such aschromia or molybdena on alumina at pressures of about 50 to 450 p.s.i.,900 to 1025 F. and at liquid hourly space velocities of about 0.1 to 10to produce a product which is fractionated into a bottoms productboiling higher than gasoline, a heavy' naphtha, a benzene fractionboiling in the range of about 160 to 185 F. and uncondensed overheadvapors. The overhead of the fractionation is stabilized at 150 to 500p.s.i. to produce a sidestream -comprising C3 and C4 hydrocarbons, alight naphtha whichis mixed with the aforesaid heavy naphtha, and astabilizer overhead -containing 25 to 75 percent by volume of hydrogen.The stabilizer overhead in part can be discharged from the systemalthough it is preferred to contact the stabilizer overhead in admixturewith the aforesaid benzene fraction with a hydrogenating catalyst athydrogenating conditions of temperature of about 300 to 800 F. andpressures of about 50 to 5000 p.s.i. The benzene fraction ishydrogenated and separated from the treated stabilizer overhead. Thehydrogenated benzene fraction is then contacted in admixture with thefeed naphtha or separately with dehydrogenating catalyst at 900 to 1025F., at 50 to 450 p.s;i., and at liquid hourly space velocities of about0.1 to 10. The patentee also statesthat for naphthas having end boilingpoints below 450 F. diphenyl may be employed and for naphthas havingboiling temperatures below 375 to 400 F. naphthalene can be used.

Since in the patented process gasoline is produced in admixture with theproduced hydrogen and dehydrogenated benzene or diphenyl or naphthalenethe condition of dehydrogenation must -be regulated to be the optimumvfor the production of gasoline of the required octane rating regardlessof the conditions which are optimum for recovery of the hydrogen and forminimum degradation of the naphthalene, diphenyl, xylene, toluene, orbenzene. Degradation of the aromatic hydrocarbons by dealkylation duringdehydrogenation dilutes the hydrogen produced and undoubtedly at leastin part accounts for the failure to exploit this process industrially.

In direct contrast to the foregoing method of returning to the reformera hydrogen donor or a mixture of dehydrogenated hydrogen donor andhydrogen the present method provides for subjecting the fat extractingagent to dehydrogenating conditions which are optimum fordehydrogenation without substantial degradation of the extracting agent.

In contrast to the off-gas from the catalytic cracking of gas oil whichcontains about 35 percent by volume of hydrogen (Table I) the make-gasof reforming unit treating Mid-Continent naphtha having a boiling rangeof about 160 to about 360 F. to produ-ce C5+ reformate having an octanerating (Research-F5 ml. TEL) of 100 has a hydrogen concentration ofabout 73 percent by volume as is shown in Table II. On the other hand,the make-gas from a reforming unit treating a Mid-Eastern naphtha havinga boiling range of about 160 to about 360 F. to produce C3+ reformatehaving an octane rating (Research-l-S ml. TEL) of 100 has a hydrogenconcentration of about 61 percent by volume as is shown in Table III.

TABLE II [Make-Gas from Mid-Continent Naphtha at 100 O.N. (Research-i-Sm1. TEL)] TABLE III Component: Mol percent Hydrogen 61.0 C1 22.0 C2 12.0C3 4.0 I-C.L 0.5 N4C3 0.4 I-C5 0.1

Total 100.0

Hydrogen sulfide, p.p.m. 20

In many reneries the off-gas of the naphtha pretreater is contacted withthe raw naphtha to remove C4 and heavier hydrocarbons from thepretreater -gas while removing water and oxygen from the raw naphtha.Hence, the concentrations of C4 and heavier concentrations are lowenough to be disregarded. In most refineries the naphtha pretreater gasfrom which the hydrogen is to be recovered has the lfollowing4composition (Table IV) TABLE IV Component: Mol percent Hydrogen 65.0 C120.0 C3 11.0 C3 3.7 C4+ about 0.2 Hydrogen sulfide 0.1

However, it is conventional because of local antipollution laws toremove the hydrogen sulfide from this gas. The hydrogen sulfide contentis generally reduced below p.p.m.

The off-gas from a hydrodecontaminating unit for the treatment ofdomestic fuel oil, kerosine, jet fuel and the like has the averagecomposition set forth in Table V.

TABLE V Component: Mol percent Hydrogen 50.0 C1 31.0 C3 14.0 C3 4.5 C40.5 Hydrogen sulfide, p.p.m. 10

In general, the aforedescribed off-gases will be treated to reduce thesulfur content to at least 10,000 p.p.m. by volume when using asulf-active hydrogenating catalyst and to at least 400 p.p.m. by volumewhen using a platinum-group metal hydrogenating catalyst. Suitableextracting agents are pure single or multi-ring aromatic hydrocarbons,pure single or multi-ring hydrocarbons having one or more saturated,i.e., alkyl side-chains, e.g., benzene, toluene, xylenes, trimethylbenzenes, ethyl benzene, diethyl benzene, ethyl-methyl benzene,naphthalene, methyl naphthalene, anthracene, methyl anthracene,phenanthrene, methyl phenanthrene, mixtures of the foregoing, aromaticpetroleum fractions such as synthetic tower bottoms obtained in thecatalytic cracking of gas oil, mixtures of aromatic hydrocarbonsrecovered from the extract obtained by extraction aromatichydrocarbon-containing naptha, kerosine synthetic tower bottoms, and, ingeneral, any aromatic hydrocarbon mixture with a selective solvent suchas sulfur dioxide, alkylene glycols, e.g., diethylene glycol,triethylene glycol, Chlorex and others. In general, the hydrogenextracting agent comprises aromatic hydrocarbon substantially devoid ofhydrogenatable material not readily dehydrogenatable at extracting agentregeneration conditions without the production of unsaturatedhydrocarbons. Suitable hydrogenation catalysts are hydrogenationcatalysts having a hydrogenation component selected from the groupconsisting of the metals having atomic numbers 44 to 96 inclusive and 76to 78 inclusive on refractory oxide base comprising alumina,silica-alumina, silica-zirconia. The

'hydrogen-ation catalyst can be selected on the basis-of the "catalystpoisons present in the gas from which hydrogen 'is `to be recovered.Thus, when the hydrogen-containing gas contains hydrogen sulfide it ispreferred to use a 'catalyst which 'is not seriousiy Vinactivated bysulfur, eg., -having la hydrogenation component such as nickeltungstensulfide, tungsten disulfide, nickel-molybdenum sulfide. -On the votherhand, when the sulfur content of the #hydrogen-containing ,gas to betreated does not exceed 30 p.p.m. A(parts per million) by weight of theacceptor platinum-group metal hydrogenating vcatalysts such as a`catalyst comprisingabou-t 0.30 to about 1.00 percent by weight ofplatinum, palladium, osmium, iridium on alu- Vmina `base can be used.However, when desirable or .necessary or convenient, the gas containingthe hydrogen :sulfide can be contacted with an aqueous solution of-al-kanolamine such -as diethanolamine or other treating agent forremoval of the hydrogen sulfide to provide a hydrogen-containing gascontaining not more than ,p.-p.=m. of sulfur by weight based on thetreating agent and hydrogenating catalyst having as the hydrogenatingcomponent platinum-group meta-l used to recover hydrogen from `theso-treated hydrogen-containing gas. n

Hydrogenating conditions are as set forth in Table VI.

TABLE v1 Gperating Conditions Broad Preferred Temperature, It.,A 200tomo 30o to 650. Pressure, p.s.i.g 150 to 2.500-... 500 to 1,500.'Liquid Hourly Space Velocity (v.'/hr./ 0.25 t0 10 0.5 to 4.

vw l

1 v. .,-volurne ol treating agent in barrels: v.-volume of hydrogenating`catalyst (barrel).

able dehydrogenating vcatalysts are catalysts comprising platinum-groupmetal on refractory oxide support. Thus, for example, a dehydrogenatingcatalyst comprising about -0.35 to about 0.6 percent `of platinum oralumina support is suitable. However, the preferred dehydrogenatingcatalyst comprises about 0.35 to about 0.6 percent yby weight ofplatinum and not more than about 0.7 percent by weight of chlorine onalumina. Y

Dehydrogenating conditions are set forth in Table VII.

4TABLE VII Operating Conditions ABroad Preferred Temperature, F v i 800to 1,050 850 to 950. Pressure, psig 100 to 1.000.... 200 to 500. LiquidHourly Space Velocity v.fea/ 0.5 to 100 2 to 30.

hr. v. Hydrogsen Circulation (mol of purified 1 to 20.V 3 to 10.

hydrogen per mol of FEA).

fea, FEA-Fat Extracting Agent.

For the recovery of hydrogen from dilute hydrogencontaining gas such as'the off-gas from catalytic cracking of gas oil, naphtha pretreateroff-gas, oit-gas from middle distillate hydrodecontaminat-ion andsimilar refinery olfgases containing at least percent by volume ofhydro- :gen the dilute `refinery gas is contacted with the extractingagent -i-n a lean extracting agent-to-hydrogen mol ratio in the range of0.20 to 5.

Illustrative of the flow of liquids and gases in the present method ofrecovery hydrogen of `at least 90 percent purity from refinery gascontaining at least 20 percent of hydrogen `are the flow sheets ofFIGURES 1 and 2.

The flow sheet FIGURE 1 illustrates the use of two hydrogen extractorsor reactors and one fat extracting agent regenerator or reactor. Thehydrogen extractors or reactors are used alternately. That is tonsay,while one is on-stream the other is off-stream and the catalyst isregenerated usually by combustion of the coke deposited on the catalystduring the on-st-ream period of the cycle in any inert gas containinglfree oxygen. The deposition of coke in the fat extract-ing agentregenerator or reactor is minimal since .temperatureand liquid hourlyspace velocity are `correlated to minimize hydro-cracking anddealkylation of alkyl-substituted extracting agent. It is -to be notedthat regeneration of the fat extracting agent FEA is an endothermicreaction which under adiabatic reaction conditions results in atemperature drop in the reactor. Since Adealkyl'ation '-ofalkyl-substituted acceptors occurs at temperatures below that at whichdehydrogenation of the 'hydrogenated extracting agent occurs it isdesirable, and in fact preferred, to separate the regenerated or leanextracting agent from the dehydrogenating catalyst substantially as soonas the temperature of the reactan-t reaches that at which thedehydrogenation reaction ceases. This is mos-t readily accomplished bycontacting the fat extracting agent with the dehydrogenating catalyst ata catalyst-to-oil ratio in the range of about 0.2 to 0.5 tons ofcatalyst per 1000 barrels of fat -extracting agent per day, i`.e., at aliquid hourly space velocity (v.fe /hr./v.c) in the range of about l0 toabout 25 when using a dehydrogenating catalyst comprising about 0.35 toabout 0.6 percent by weight of platinum and `not more than 0.7 percentof chlorine and/ or fluorine on alumina support. y

As illustrated in FIGURE 1 at start-up a lean extracting agent such asthe hydrocarbons recovered from the extraction of kerosine with sulfurdioxide or the synthetic tower bottoms recovered from the product of thecatalytic cracking of gas oil, or in general a hydrogenatabile aromatichydrocarbon comprising at least 60 volume percent including volumepercent, of aromatic hydrocarbon and substantially devoid of otherhydrogenatable material is drawn from a source not shown through pipe 1under contnol of val-ve 2. After start-up only make-up quantities offresh extracting agent are drawn through valve 2. The extracting agent(lean) ilows through pipe 1 to the suction side of pump 3. Pump 3discharges the `lean extracting agent (LEA) into pipe 4 at a pressure inexcess of that in the on-stream extractor or reactor 15 or 16 by theamount of pressure drop between pump 3 and the 'on-stream extractor orreact-or. The lean extracting agent ows through pipe 4 to indirect heatexchanger 5 where the lean extracting agent is in heat transfer relationwith lthe eiiluent of the on-s'tream extractor or reactor flowing fromeiuent manifold 26. From heat exchanger 5 the lean extracting agent owsthrough pipe 6 to coil 7 in heater or furnace 8. In furnace 8 Vthe leanextracting agent is heated to a temperature to `at least 2.0 percent byvolume of hydrogen contains not more than about 10,000 p.p.m. by volumeof sulfur as hydrogen suliide.)

Contemporaneously with the flow of lean extracting agen-trthroughreactor 15 refinery off-gas containing at least 20 percent by volume ofhydrogen and not more `than 10,000 p.p.m. by volume of sulfur ashydrogen suliide ows from a source not yshown through conduit lene.

' much as live.

17 to compressor 18. Compressor 18 discharges the refinery off-gas intoconduit 19 at a pres-sure in excess of that in reactor 15 by thepressure drop b-etween cornpressor 18 and reactor 15. The compressedrefinery offgas flows through conduit 19 to coil 20 in furnace or heater21. In heater 21 the refinery off-gas is heated to a temperature suchthat when mixed with the heated lean extracting agent in inlet manifold10 the mixture will have a hydrogenating temperature. The heated leanextracting agent and heated refinery off-gas flow respectively throughpipe 9 and conduit 22 to inlet manifold 10.

The lean extracting agent-to-hydrogen mol ratio depends on theextracting agent employed. For example, if naphthalene is used, one molcan bind as many as five mols of hydrogen when completely hydrogenatedin accordance with the following equation:

CioHs-i-'ZHz'e-CioHiz-t 3H2C1nHis Naphthalene Tetralin Decalin Thus, thelean extracting agent-to-hydrogen mol ratio is 0.20 (l/) for completehydrogenation of the naphtha- For partial hydrogenation (to tetralin),the lean extracting agent-to-hydrogen mol ratio is 0.50 (l/ 2). In thecase of benzene, one mol can react with a total of three mols ofhydrogen when completely hydrogenated in accordance with the followingequation:

The lean extracting agent-to-hydrogen mol ratio for benzene is 0.33(1/3). H-owever, to ensure maximum recovery of hydrogen from the dilute(with respect to hydrogen) refinery olf-gas an excess of hydrogenatablearomatic hydrocarbon above that which theoretically can be hydrogenatedby the hydrogen of the dilute refinery oE-gas is generally used.Accordingly, it is preferred to employ lean extracting agent-to-hydrogenmol ratios as This excess of lean extracting agent not only ensuresmaximum extraction of the hydrogen from the dilute refinery off-gas butalso more ready control of the temperature of the endothermic reaction.

Returning vto a description of FIGURE 1; the heated lean extractingagent and dilute refinery off-gas are mixed in inlet manifold 10 in theproportion of at least 0.20 mol of hydrogenatable aromatic hydrocarbonper mol of hydrogen in the dilute refinery off-gas to provide anextraction mixture. The extraction mixture at hydrogenation temperatureflows through inlet manifold 10 to manifold branch 11 (valve 14 closed;valve 12 open). The extraction mixture flows downwardly in recator 15 incontact with hydrogenation catalyst. The fat extracting agent, i.e., atleast partially hydrogenated aromatic hydrocarbon, and denuded refineryoff-gas, designated extractor effluent, flow from reactor 15 throughmanifold branch 23 (valve 24 open; valve 28 closed) to outlet manifold26. The extractor effluent flows through outlet manifold 26 to indirectheat exchanger 5 to conduit 30 and cooler 31. In cooler 31 the extractoreffluent is cooled to a temperature at which the fat extracting agen-tis liquid at the existing pressure. The cooled extractor efuent flowsfrom cooledl 31 through conduit 32 to liquid-gas separator 33. Inliquid-gas separator 33 the uncondensed denuded refinery off-gasseparates from the condensed fat extracting agent. The uncondenseddenuded refinery off-gas comprising C1 to C5 hydrocarbons flows fromseparator 33 through conduit 34 to hydrocarbon recovery, processing, therefinery fuel main, or the refinery flare.

The condensed fat extracting agent flows from separator 33 through pipe35 to the suction side of pump 36. The fat extracting .agent isdischarged into pipe 37 by pump 36 at a pressure greater than thepressure in reactor 25 by the pressure drop between pump 36 and reactor25.

The fat extracting agent flows through pipe 37 to indirect heatexchanger 38. In indirect heat exchanger 38 the fat extracting agent isin heat transfer relation with the effluent of reactor 25 comprisinghydrogen and lean extracting agent flowing from reactor 25 throughconduit 43. From indirect heat exchanger 38 the fat extracting agentflows through conduit 39 to coil 40 in furnace or heater 41. At a pointin conduit 39, intermediate to indirect heat exchanger 38 and to furnace41, percent or better hydrogen flowing from compressor 51 throughconduit 52 is mixed with the fat extracting agent in the mol ratio ofaboutl 1 to about 20, preferably about 3 to 10 mols of hydrogen per molof hydrogenated aromatic hydrocarbon -to provide a recovery charge.

The recovery charge is heated in furnace 41 to a dehydrogenatingtemperature to provide a vapor inlet temperature at reactor 25 in therange of aboutSOO" to about 1050" F. The heated recovery charge flowsfrom heater 41 through conduit 42 to reactor 25. The recovery chargeflows downwardly through reactor 25. The extracting agent regeneratoreluent comprising hydrogen and lean extracting agent flows from reactor25 through conduit 43, indirect heat exchange 38 and conduit 44 tocooler 45. In cooler 45 the extracting agent regenerator effluent iscooled to a temperature at which the extracting agent is condensed atthe existing pressure. The uncondensed hydrogen and minor concentrationof hydrocarbons together with the condensed lean extracting agent flowfrom cooler 45 through conduit 46 to liquid-gas separator 47.

In liquid-gas separator 47 the uncondensed recovered hydrogen andhydrocarbons boiling below the boiling point of the extracting agentseparate from the condensed lean extracting agent and ow therefrom, asla gaseous stream containing at least 90 percent by volume of hydrogenbalance, to make percent by volume, hydrocarbons boiling below theboiling point of the extracting agent, through conduit 48 tohydroprocesses. A portion of the at least 90 percent pure hydrogen isdivertcd from conduit 48 under control of valve 50 to conduit 49 and thesuction side of compressor 51. The quantity of 90 percent pure hydrogenso diverted is sufficient to provide the aforesaid 1 to 20, preferably.3 to 10 mols of hydrogen per mol of hydrogenated aromatic hydrocarbonin the recovery charge. Compressor 51 recompresses the 90 percent purehydrogen to a pressure in excess of that in conduit 39. The recompressed90 percent pure hydrogen flows from compressor 51 through conduit 52 toconduit 39.

The condensed lean extracting agent (LEA) separated in liquid-gasseparator 47 from the uncon-densed at least 90 percent pure hydrogen,flows from separator 47 through pipe 53 and 1 to the suction side ofpump 3 (valve 54 open; valve 2 closed except to admit make-up extractingagent to pipe 1). The lean extracting agent is then recycled to thehydrogenating-reactor or extractor 1S or 16 which is on-stream toextract hydrogen from further quantities of refinery off-gas containingat least 20 percent by volume of hydrogen. Those skilled in the art willrecognize that for simplicity the piping required for regeneration ofcoked catalyst has been omitted since regentration of catalystdeactivated with coke is too well known to require description. Briefly,the catalyst in the off-stream extractor, presently assumed fordescriptive purposes to be reactor 16, is regenerated by isolating thereactor or extractor from the extracting agent and refinery off-gas andfrom extractor effluent manifold. The extractor or reactor is pluggedVwith inert gas such as flue gas, steam, nitrogen. The purged reactor isthen pressured with flue gas and the flue gas heated and encirculateduntil the temperature in the catalyst bed is at a burning temperature.Free oxygen, e.g., air is then admixed which the circulating inert gasand the coke burned from the catalyst. When the burning front reachesthe bottom of the catalyst bed and free oxygen is detected in theeffluent gas the supply of air is cut-off and the reactor purged untilthe `oxygen content of the etliuent gases does not exceed 1 mol-percent.The reactor or extractor is then pressured with refinery off-gas to beextracted and the ilow of heated lead extracting agent and reneryoff-gas to be extracted diverted from the on-steam extractor -or reactorto the olf-steam extractor in which the catalyst has been regenerated'.

FIGURE 2 is a flow sheet likewise illustrating the use of three:reactors for extracting hydrogen from dilute reinery gas streams andrecovering they extracted hydrogen. However, whereasv inthe embodimentillustrated in FIG- URE l one reactor is used only for the recovery ofextracted hydrogen while the. other two reactors are used alternativelyfor extracting. hydrogen in the embodiment illustrated in FIGURE 2 allof the reactors are charged with a catalyst'y having not onlyhydrogenating capabilities but also having dehydrogenating capabilities.

Catalysts having hydrogenating and dehydrogenating capabilities suitablefor use in the present method are for example, a mixture of oxides ofcobalt and molybdenum on alumina support; platinum-group metal onrefractory oxide support, such as` alumina, silica-alumina, and thelike; and catalysts comprising a mixture of at least 18 percent byweight of at least one. oxide of chrominum, molybdenum, and vanadium,and at least one refractory oxide such as alumina, silica, zirconia.

i The embodiment illustrated in FIGURE 2 has parvticular advantages whenrecovering hydrogen from dilute refinery gas containing more than 30p.p.in. of sulfur 'based upon the weight of the hydrogenatable aromatichydrocarbons used as, for example, when recovering hydrogen from thehydrogen-containing off-gases from a hydrodesulfurizing unit treatingmiddle distillate fuels such as kerosine, domestic fuel oil, jet fuelsand the like. For example, the hydrogen-containing gas from ahydrodesulfurization unit hydfrotreating a middle distillate e.g.,domesticV fuel oil has the following average cornposition:

t P.p.ni.-parts per million by volume on gas. While a dilute off-gascontaining such a concentration of sulfur and ammonia must be pretreatedto reduce the sulfur concentration to not more than 30 p.p.m. by weightbased on the extracting agent (15 to 400 ppm. by volume :based on thegas for treating agent-to-hydrogen mol ratios in the range of 0.20 to 5)and the ammonia concentration to not more than 5 p.p.m. by weight basedon the said extracting agent (5 Ito 150 p.p.in. by volume based on thegas for hydrogenatable aromatic-to-hydrogen mol ratios in the range of0.20 to 5 when platinum-group metal catalyst having hydrogenating anddehydrogenating capabilities is employed for extracting hydrogen (fromdilute off-gas) and recovering hydrogen (from lean extracting agent), incontract when using sulfur and nitro- -gen-insensitive catalyst havinghydrogenating and dehydrogenating capabilities such as a mixture ofoxides of chromium and aluminum, or of oxides of chromium, molybdenum,and aluminum, or of oxides of cobalt,

lmolybdenum, and aluminum, the concentrations of sulfur and ammonia inthe dilute off-gas can be as high as 10,000 p.p.m. by volume and 1500p.p.m. by volume respectively, i.e., 750 p.p.ni. and 50 p.p.in. byweight based on the lean extracting agent using a mol ratio in the rangeof 0.20 to 5 mols of said aromatic hydrocarbon per mol of hydrogen.Dilute off-gas such as the aforesaid dilute off-gas from ahydrodesulfurization unit trea-ting middle 10 distillate fuel oil ascharacterized in Table VIII is-illustrative of dilute off-gas treated inaccordance with the flow-sheet of FIGURE 2.

Reactors 114, 160, and 141 are charged with sulfur andnitrogen-insensitive catalyst having both hydrogenating anddehydrogenating capabilities. Lean extracting agent, for example,synthetic crude. tower bottoms from a catalytic cracking unit having theaverage composition set forth in Tabley IX is drawn from a source notshown through pipe 101 under control of valve. 102 by pump 103.

TABLE` IX [Lean Extracting Agent-#TCC Synthetic Crude Tower- Bottoms]Boiling Range, F. 300 to 371 IBP, F. 300 5%, F. 314 10%, F. 3-30 F. 344EBP, F. 37'1 Aromatics (principally polyalkylbenzenes) 60-61`.8 vol.percent.

The lean extracting agent is discharged by pump 103 into pipe 104 at apressure greater than that in the reactor of reactors 114, 160, and 141on-stream by at least the pressure drop between pump 103 and theon-stream reactor. The lean extracting agent flows Ithrough pipe 104 tocoil 105 in furnace or heater 106. In heater 106 the lean extractingagent is heated to a temperature such that when mixed with diluteoif-gas in the ratio of about 0.20 to about 5 mols of hydrogenatablearomatic per mol of hydrogen in the dilute oif-gas the extractionmixture soformed has a hydrogenating temperature at the vapor inlet ofthe on-stream extractor or hydrogenating reactor. (For simplicity itwill be assumed that the following cycle having the indicated phases isemployed.)

TABLE X Ri (114) R2 (160) Ra (141) First Phase Extracting IdleRegeneration of Hydrogen. Extracting Agent,

Second Phase Regeneration of Regeneration of Extraction of Catalyst.Extracting Hydrogen to Agent. about 10% coke on catalyst.

Third Phase Regeneration of 'Extraction to Regeneration of extractingabout 10% catalyst. agent. coke on catalyst.

Fourth Phase Extraction to Regeneration of Regeneration of about 10%catalyst. extracting coke on cataagent.

` lyst.

Fifth Phase Regeneration of Regeneration of Extraction to catalyst.extracting about 10% coke agent. on catalyst.

Sixth Phase Regeneration of Extraction to Regeneration of extractingabout 10%. catalyst. agent. coke on catalyst.

The heated lean extracting agent Hows from furnace 106 through conduit107 to inlet manifold 108 and manifold branch 112 (valve 113 open;valves 159 and 140 closed). At a point in conduit 107 intermediate toinlet manifold 10S and heater 106 the dilute oif-gas from which hydrogenis to be recovered is mixed with the lean extracting agent in a molratio in the range of 0.20 to 5 mols of hydrogenatable aromatic per molof hydrogen. The dilute off-gas iiows from a source not shown throughconduit 110 to the suction side of compressor 109. Compressor 109compresses the off-gas to a pressure at least equal to that in conduit107 and discharges the off-gas into conduit 111 through which theoff-gas ows to conduit 107 and admixture with the lean extracting agent.(Those skilled in the art will appreciate that when the off-gas isavailable contiguous to conduit 107 at a pressure at least equal to thatin conduit 107 recompression is unnecessary and compressor 109 can bebypassed.)

y, The extractor mixture comprising off-gas and lean extracting agent inproportions to provide 0.20 to mols `of hydrogenatable aromatic per molof hydrogen in the olf-gas flows through conduit 107 to extractor inletmanifold 108. From extractor inlet manifold 108 in the first yphase ofthe cycle the extractor mixture ows through manifold branch 112 (valve113 open; valves 159 and 140 closed; valves 166, and 157 closed) toextractor reactor 114. The extractor mixture flows downwardly inextractor 114 in contact with particle-form solid hydrogenating catalyst'in this description sulfurand nitrogeninsensitive hydrogenatingcatalyst. T he extractor effluent comprising fat extracting agent andextracted off-gas flows from extractor 114 to manifold branch 11.5.(Valve 116 open; valves 168, 162, 164, and 143 closed.) The extractoretiluent flows through manifold branch 115 to extractor elfluentmanifold 117. The extractor effluent flows through extractor efuentmanifold 117 to indirect heat exchanger 118 Where the extractor effluentis in heat transfer relation with lean extracting agent flowing fromindirect heat exchanger 120 through conduit 154. From indirect heatexchanger 118 the extractor efliuent ows through conduit 119 to indirectheat exchanger 120 where the extractor efuent is in heat transferrelation with lean extracting agent flowing from liquid-gas separator151 through pipe 153. From indirect heat exchanger 120 the extractoreiuent flows through conduit 121 to cooler 122. In cooler 122 theextractor euent is cooled to a temperature at which the fat extractingagent is condensed at the existing pressure. (Fat extractving agent asused hereinbefore, here, and hereinafter designates extracting agent inwhich the hydrogenatable arolmatic has been in contact withhydrogen-containing gas in the presence of hydrogenating catalyst underhydrogenating conditions of temperature, pressure, and liquid hourlyspace velocity regardless of the degree of completeness ofthehydrogenation of said hydrogenatable aromatic hydrocarbon. Thus, forextracting agents containing hydrogenatable aromatic hydrocarbons suchas benzene and alkylbenzenes it is immaterial whether all of themolecules each have accepted six atoms of hydrogen or only one atom ofhydrogen or that a portion of the molecules of the aromatic each haveaccepted one atom of hydrogen and the balance none of a portion of themolecules of the aromatic each have accepted one to six atoms ofhydrogen and the balance none.) The cooled extractor etiluent comprisingextracted off-gas and fat extracting agent flows from cooler 122 throughconduit 123 to gas-liquid separator 124. In gas-liquid separator 124 theextracted off-gas separates from condensed fat extracting agent. Theextracted off-gas, designated waste gas, ows from separator 124 throughconduit 125 to means for recovery of hydrocarbons, isomerization ofhydrocarbons, hydrocarbon conversion as of methane to acetylene, therefinery Vfuel main, the refinery fiare or other means of treating ahydrocarbon stream of this composition.

The condensed fat extracting agent ows from separator 124 through pipe126 to indirect heat exchanger 127 Where the fat extracting agent is inheat transfer relation with extracting agent regenerator effluentflowing from regenerator outlet manifold 146, indirect heat exchanger129, conduitl 147 to indirect heat exchanger 127. From indirect heatexchanger 127 the fat extracting agent flows through pipe 128 toindirect heat exchanger 129 where the fat extracting agent is in heattransfer relation with the regenerator efuent flowing from fatextracting agent regerenator manifold 146. From indirect heat exchanger129 the fat extracting agent flows through pipe 130 to ythe suction sideof pump 131. Pump 131 discharges the fat extracting agent into pipe 132through which the fat extracting agent flows to coil 133 in furnace orheater 134:

In furnace 134 the fat extracting agent is heated to a temperature suchthat, when mixed with hydrogen in regenerator inlet manifold 136 toprovide a regeneration feed comprising about l to about 20 mols ofhydrogen of at least percent purity per mol of hydrogenated aromatichydrocarbon, a vapor inlet dehydrogenation temperature of about 800 toabout 1050 F. at thel regenerator-reactor inlet (141) is provided. Fromheater 134 the heated fat extracting agent (FEA) flows through pipe 135to regenerator inlet manifold 136. In regenerator inlet mani-fold 136the fat extracting agent is mixed with hydrogen of 90 percent purityflowing either from a source not shown through conduits 171 and 169under control of valve 172 or after start-up from liquid-gas separator151, and conduits 152 and 169 under control of valve 170. (Whenrecovered hydrogen or extraneous hydrogenis not available at a pressureat least equal to that in regenerator manifold 136 the extraneoushydrogen and the recovered hydrogen is compressed to the neces.- sarypressure for introduction into regenerator manifold 136.)

From regenerator inlet manifold 136 the regenerator charge mixturecomprising fat extracting agent and hydrogen in the ratio set forthhereinbefore flows through manifold branch 137 (valve 138 open; valves157 and 166 closed; valve 140 closed) and manifold branch 139 toregenerator 141. The regenerator charge mixture flows downwardly inregenerator 141 in Contact with dehydrogenating catalyst describedhereinbefore. The regenerator efuent comprising recycle hydrogen, i.e.,hydrogen admixed in regenerator manifold 136, make-hydrogen, i.e.,hydrogen removed lfromthe fat extracting agent and lean extracting agentflows from regenerator 141 through manifold branch 144 (Valve 145 open;valves 143, 164, and 168 closed) to regenerator manifold 146. Fromregenerator manifold 146 the regenerator effluent llows to indirect heatexchanger 129, conduit 147, indirect heat exchanger 127 and conduit 148to cooler 149.

In cooler 149 the regenerator effluent comprising hydrogen and leanextracting agent is cooled to a temperature at which the lean extractingagent is liquid. The condensed lean extracting agent and uncondensedhydrogen ow from cooler 149 through conduit 150 to liquid-gas separator151.

In liquid-gas separator 151 hydrogen of at least 90 percent purityseparates from lean extracting agent. The hydrogen of at least 90percent purity flows from separator 151 through conduit 152 to use inother hydro processes. A portion to -provide about 1 to about 20 mols ofhydrogen per mol of hydrogenatable aromatic hydrocarbon in regeneratormanifold 136 as described hereinbefore is diverted through conduit 169under control of valve 170 (valve 172 closed).

The condensed lean extracting agent flows from separator 151 throughpipe 153 to indirect heat exchanger 120, pipe 154, indirect heatexchanger 118 and pipe 155 to pipe 101 for use as extracting agent forthe extraction of hydrogen from further amounts of 4dilute reneryoff-gas.

As set forth in Table X supra in the first phase of a cycle reactor 114is in the extraction operation, reactor 141 is in `the regeneration ofthe extracting agent, and reactor is idle. When the activi-ty of `thecatalyst in extractor 114 is reduced to an industrially appreciableextent as indicated by a decrease -in the difference in temperaturebetween the temperature of the extractor effluent and the vapor inlet ofthe extractor, the extractor is taken .olf-stream and the catalystregenerated. Regenerator 141 is then preferably taken off regeneratingof the extracting agent and used for extracting hydrogen from the diluterefinery off-gas. Reactor 160 which in the rst phase of the cycle isidle is then placed in the ifat extracting agent regenerating operation.Operation in the second phase of the cycle then proceeds until thecatalyst in reactor 141 has been contaminated by the deposition of anamount of coke equal to about 10 percent by weight 13 of the catalysttherein as determined, for example, bythe decrease in the differencebetween the temperature of the extractor effluent and the vapor inlettemperature. (Hydrogenation being an exotherrnic reaction thetemperature .of the extractor effluent is higher than the vapor inlettemperature. The more vigorous the hydrogenation, Le., the more acti-vethe hydrogenat-ion, the `greater the temperature rise.) Reactor 141 isthen taken off-stream and the catalyst therein regenerated while reactor1160 (extracting agent regeneration in the second phase) is puton-stream as an extractor .until about percent by weight of coke isdeposited thereon and reactor 114 (catalyst now regenerated) is put.on-stream for regeneration of the extracting agent in the third phaseof the cycle. fIn the fourth phase of the cycle the .catalyst in reactor160 (extraction stage in lthe third phaser) is regenerated, reactor.1-41 (regeneration of .catalyst -in the third phase) is yused forregeneration -of the extracting agent, and reactor 114 previously usedin the regeneration of extracting agent) is used Afor extractinghydrogen from dilute refinery ott-gas until the amount of coke depositedon the catalyst amounts to about 10 percent -by weight thereof.

In the .fifth phase of .the cycle the catalyst -in reactor 114 isregenerated, reactor 1411 is used as .a hydrogen extractor, and reactor160 is used to regenerate the fat extracting agent. In the sixth phaseof the iirst cycle reactor 11'4 is .used for the regeneration of the:fat extracting agent, reactor 160 is used as an extractor of hydrogenfrom dilute refinery off-gas and the catalyst in reactor 14.1 isregenerated. This ends the first cycle.

Those skilled in the art Iwill rec-ognize that there is nothingimmutable in .the sequence in which the reactors appear in any cycle.Tha-t is to say, in the iirst phase of the `first cycle reactor 141 canbe used .as an extractor, reactor 160 used for regeneration of theextracting agent and reactor L14 can be idle. The other possiblecombination can .also be used in the first phase. However, it ispreferred after the .first phase to use the reactor in 'which thecatalyst is fresh or regenerated for regeneration of the extractingagent and to use the catalyst which in the previous phase was used `forregeneration of the extracting .agent for the extraction of hydrogenfrom refinery gas streams dilute with respect to hydrogen.

I claim:

I1. A method of recovering hydrogen from refinery gas dilute withrespect to hydrogen which comprises in a gas extractor contacting renerygas comprising at least about Z0 percent by volume of hydrogen and anextracting agent having as its sole hydrogenatable material at least onearomatic hydrocarbon with hydrogenating catalyst at hydrogenatingconditions of temperature, pressure, and liquid hourly space velocity,obtaining extracltor effluent comprising fat extracting agent andextracted refinery gas, cooling said extractor eiuent 4to a -temperatureat which said fat extracting agent is condensed at the existingpressure, separating said cooled extractor effluent into waste gascomprising refinery gas having reduced hydrogen content and fatextracting agent, in an extracting agent regeuerator contacting onlysaid separated fat extracting .agent in admixture .with hydrogen of atleast 90 percent purity with dehydrogenating catalyst at dehydrogenatingconditions of temperature, pressure, and liquid hourly space velocitywhilst maintaining a concentration of at least `90 percent .of hydrogenin the fraction of reaction vapors boiling below the boiling point ofsaid lean extracting agent, obtaining regenerator effluent comprisinghydrogen produced in said extracting agent regenerator, the .aforesaidadmixed hydrogen, and lean extracting agent, cooling said regeneratoreffluent to a temperature at which said lean extracting agent iscondensed at the existing pressure, separating said cooled regeneratoreuent into a gaseous fraction comprising at least 90 percent hydrogen byvolume and condensed lean extracting agent, recycling a portion of saidgaseous fraction to said .extracting agent regenerator andrecovering the'balance of said gaseous fraction as `hydrogen of at least 9.0 percentpurity, and recycling 'at least a portion of said condensed -leanextracting agent to the aforesaid Vgas extractor.

2. The method of claim d wherein the hydrogenating catalyst and thedehydrogenating catalyst yare of .substan- .tially the same composition.

3. The method set forth. i-n claim I1 lwherein the hydrogenatingcatalyst and the .dehydrogenating catalyst are of substantiallydifferent composition.

4. The method set forth in claim -1 wherein the hydro-.genatingcomponent of the hydrogenating catalyst is substantially thelsame as the dehydrogenating component of .the dehydrogenating catalyst.

5. The method. set forth `in .claim 1 wherein the hydrogenatingcomponent ofthe hydrogenating catalyst is different .to thedehydrogenating component of the dehydrogenating catalyst.

f6. The method set forth in claim 1 lwherein three reactors are chargedwith substantially the same quantity of particle-form solid catalyticmaterial having hydrogenating and dehydrogenating capabilities, whereina cyclic manner catalyst having a coke deposit of about 10 percent byweight based on the catalyst is regenerated, then used -for extractinghydrogen from refinery gas dilu-te with respect to hydrogen.

7. The method set forth in claim 6 wherein the particle-form solidcatalytic material is sulfur and nitrogen insensitive.

y8. The method set forth .in claim 6 wherein the particle-form solidcatalytic material is platinum-group metal and the refinery gas to beextracted contains not more .than 400 p.p.m. volume of hydrogen sulfide.

9. A method of recovering hydrogen from refinery gas dilute with respectAto hydrogen which comprises in a gas extractor contacting renery gascomprising at least about 20 percent by volume of hydrogen and anextracting agent having as its sole hydrogenatable material at least onearomatic hydrocarbon with hydrogenating catalyst at hydrogenatingconditions of temperature, pressure, and liquid hourly space velocity,obtaining extractor effluent comprising fat extracting agent andextracted renery gas, cooling said extractor effluent to a temperatureat which said fat extracting agent is condensed at the existingpressure, separating said cooled extractor effluent into waste gascomprising .refinery gas having reduced -hydrogen content and fatextracting agent, in an extracting agent regenerator contacting onlysaid separated fat extracting agent admixed lwith hydrogen of at least Apercent purity with platinum-group metal dehydrogenating catalyst atdehydrogenating conditions of temperature, -pressure and liquid hourlyspace velocity 'whilst maintaining a concentration of at least 90percent of hydrogen in the .fraction of reaction vapors boiling belowthe boiling point of said lean extracting agent, separating regeneratorvapor from dehydrogenating catalyst after said fat extracting agent hascontacted not more than 0.2 to 0.5 ton of said platinum-group metaldehydrogenating catalyst per 1000 barrels of extracting agentregenerated per day, said separated regenerator vapors comprising saidadmixed hydrogen, hydrogen produced in said regenerator, and leanextracting agent, cooling said separated regenerator vapors to atemperature at which said lean extracting agent is condensed at theexisting pressure, separating said cooled regenerator vapors into agaseous fraction containing at least 90 percent by volume of hydrogenand liquid lean extracting agent, recycling a portion of said gaseousfraction to the regenerator to supply the admixed hydrogen of at .least90 percent purity, recovering the balance of said gaseous fraction ashydrogen of at least 90 percent puri-ty, and recycling at least aportion of said lean extracting agent to said extractor.

10; The method set forth in claim 9 wherein the hy- -drogenatingcatalyst is selected from the group consisting of mixture v,of oxides ofcobalt and molybdenum,

knickel-tungsten, sul-tide, tungsten disulde, and. nickelmolybdenumsulde..

11. The method set forth in claim 9 wherein the dehydrogenating catalystcomprises about 0.35 to about 0.6 percent by weigh-t of platinum onalumina support.

i finery gas is treated to reduce the sulfur and nitrogen con-- tentthereof to not more than 400 p.p.m. by volume and 150 p.p.m. by volumerespectively, wherein the hydrogenating catalyst comprises about 0.35 toabout 0.6 percent by weight of platinum on alumina and wherein the lplatinum group metal dehydrogenatng catalyst comprises about 0.35 toabout 0.6 percent by Weight of platinum on alumina support.

14. The method set forth in claim 9 wherein the renery gas contains notmore than 10,000 p.p.m. by volu-me of sulfur and not more than 1,500p.p.m. by volume of nitrogen, wherein the hydrogenating-catalyst isnickeltungstein sulfide, and wherein the dehydrogenating catalystcomprises about 0.35 to about 0.6 percent by weight 0f platinum onalumina support.

References Cit-ed by the Examiner UNITED STATES PATENTS 2,328,828 9/1943Marschner 208-516 2,502,958 4195O Johnson 208-56 X 2,913,401 11/1959Weikart et al 23-212 X MAURICE A. BRINDISI, Primary Examiner.

1. A METHOD OF RECOVERING HYDROGEN FROM REFINERY GAS DILUTE WITH RESPECTTO HYDROGEN WHICH COMPRISES IN A GAS EXTRACTOR CONTACTING REFINERY GASCOMPRISING AT LEAST ABOUT 20 PERCENT BY VOLUME OF HYDROGEN AND ANEXTRACTING AGENT HAVING AS ITS SOLE HYDROGENATABLE MATERIAL AT LEAST ONEAROMATIC HYDROCARBON WITH HYDROGENATING CATALYST AT HYDROGENATINGCONDITIONS OF TEMPERATURE, PRESSURE, AND LIQUID HOURLY SPACE VELOCITY,OBTAINING EXTRACTOR EFFLUENT COMPRISING FAT EXTRACTING AGENT ANDEXTRACTED REFINERY GAS, COOLING SAID EXTRACTOR EFFLUENT TO A TEMPERATUREAT WHICH SAID FAT EXTRACTING AGENT IS CONDENSED AT THE EXISTINGPRESSURE, SEPARATING SAID COOLED EXTRACTOR EFFLUENT INTO WASTE GASCOMPRISING REFINERY GAS HAVING REDUCED HYDROGEN CONTENT AND FATEXTRACTING AGENT, IN A EXTRACTING AGENT REGENERATOR CONTACTING ONLY SAIDSEPARATED FAT EXTRACTING AGENT IN ADMIXTURE WITH HYDROGEN OF AT LEAST 90PERCENT PURITY WITH DEHYDROGENATING CATALYST AT DEHYDROGENATINGCONDITIONS OF TEMPERATURE, PRESURE, AND LIQUID HOURLY SPACE VELOCITYWHILST MAINTAINING A CONCENTRATION OF AT LEAST 90 PERCENT OF HYDROGEN INTHE FRACTION OF REACTION VAPORS VOILING BELOW THE BOILING POINT OF SAIDLEAN EXTRACTING AGENT, OBTAINING REGENERATOR EFFLUENT COMPRISINGHYDROGEN PRODUCED IN SAID EXTRACTING AGENT REGENEREATOR, THE AFORESAIDADMIXED HYDROGEN, AND LEAN EXTRACTING AGENT, COOLING SAID REGENERATOREFFLUENT TO A TEMPERATURE AT WHICH SAID LEAN EXTRACTING AGENT ISCONDENSED AT THE EXISTING PRESSURE, SEPARATING SAID COOLED REGENERATOREFFLUENT INTO A GASEOUS FRACTION COMPRISING AT LEAST 90 PERCENT HYDROGENBY VOLUME AND CONDENSED LEAN EXTRACTING AGENT, RECYCLING A PORTION OFSAID GASEOUS FRACTION TO SAID EXTRACTING AGENT REGENERATOR ANDRECOVERING THE BALANCE OF SAID GASEOUS FRACTION AS HYDROGEN OF AT LEAST90 PERCENT PURITY, AND RECYCLING AT LEAST A PORTION OF SAID CONCENSEDLEAN EXTRACTING AGENT TO THE AFORESAID GAS EXTRACTOR.